Continuous and batch cultures of Escherichia coli KJ134

van Heerden and Nicol Microbial Cell Factories 2013, 12:80http://www.microbialcellfactories.com/content/12/1/80

RESEARCH Open Access

Continuous and batch cultures of Escherichia coliKJ134 for succinic acid fermentation: metabolicflux distributions and production characteristicsCarel D van Heerden† and Willie Nicol*†

Abstract

Background: Succinic acid (SA) has become a prominent biobased platform chemical with global productionquantities increasing annually. Numerous genetically modified E. coli strains have been developed with the mainaim of increasing the SA yield of the organic carbon source. In this study, a promising SA-producing strain, E. coliKJ134 [Biotechnol. Bioeng. 101:881–893, 2008], from the Department of Microbiology and Cell Science of theUniversity of Florida was evaluated under continuous and batch conditions using D-glucose and CO2 in a mineralsalt medium. Production characteristics entailing growth and maintenance rates, growth termination points andmetabolic flux distributions under growth and non-growth conditions were determined.

Results: The culture remained stable for weeks under continuous conditions. Under growth conditions the redoxrequirements of the reductive tricarboxylic acid (TCA) cycle was solely balanced by acetic acid (AcA) production viathe pyruvate dehydrogenase route resulting in a molar ratio of SA:AcA of two. A maximum growth rate of 0.22 h-1

was obtained, while complete growth inhibition occurred at a SA concentration of 18 g L-1. Batch culture revealedthat high-yield succinate production (via oxidative TCA or glyoxylate redox balancing) occurred under non-growthconditions where a SA:AcA molar ratio of up to five was attained, with a final SA yield of 0.94 g g-1. Growthtermination of the batch culture was in agreement with that of the continuous culture. The maximum maintenanceproduction rate of SA under batch conditions was found to be 0.6 g g-1 h-1. This is twice the maintenance rateobserved in the continuous runs.

Conclusions: The study revealed that the metabolic flux of E. coli KJ134 differs significantly for growth andnon-growth conditions, with non-growth conditions resulting in higher SA:AcA ratios and SA yields. Bioreactioncharacteristics entailing growth and maintenance rates, as well as growth termination markers will guide futurefermentor designs and improvements.

Keywords: Escherichia coli KJ134, Succinic acid, Metabolic flux distribution, Chemostat, Growth inhibition,Non-growth production, D-glucose, Mineral salt medium

BackgroundSuccinic acid (SA) is poised to become a major bulk-scalebiobased chemical. The projected annual commercial pro-duction of ‘biosuccinic acid’ at the end of 2015 is estimatedto be in excess of 150 ktons per annum, with five separatecompanies (BASF-Purac, Bioamber, Reverdia (DSM-Ro-quette), Mitshubishi-PPT and Myriant) currently develo-ping and implementing pilot- and industrial-scale plants

* Correspondence: [email protected]†Equal contributorsDepartment of Chemical Engineering, University of Pretoria, Hatfield, Pretoria0002, South Africa

© 2013 van Heerden and Nicol; licensee BioMCreative Commons Attribution License (http:/distribution, and reproduction in any medium

[1,2]. Although traditionally seen as a speciality chemical,new applications like the production of polybutylene suc-cinate to manufacture biodegradable plastic have stimu-lated research on, and the development of, fermentativeSA production. SA can also replace petrochemical buildingblocks to produce intermediate chemicals like tetrahydro-furan, γ-butyrolactone, 1,4-butanediol and maleic anhyd-ride. Various reviews are available on the general topic ofproducing SA with biocatalysts [3-9]. Additionally, reviewsby Beauprez et al. [10] and Cheng et al. [11] focus specific-ally on genetically modified microorganisms. Cukalovicet al. [7] focus on the viability of industrial biochemical SA

ed Central Ltd. This is an Open Access article distributed under the terms of the/creativecommons.org/licenses/by/2.0), which permits unrestricted use,, provided the original work is properly cited.

mailto:[email protected]

http://creativecommons.org/licenses/by/2.0

Table 1 Gene modifications to E. coli C to obtain thesuccinate-producing E. coli KJ134 [14,24]

Enzyme Modification Abbreviation

2-ketobutyrate formate lyase Inactivation Δ tdcE

Acetate kinase Inactivation Δ ackA

Alcohol dehydrogenase Inactivation Δ adhE

Aspartate aminotransferase Inactivation Δ aspC

Citrate lyase Inactivation Δ citF

Formate transporter Inactivation Δ focA

Lactate dehydrogenase Inactivation Δ ldhA

Methylglyoxal synthase Inactivation ΔmgsA

NAD+-linked malic enzyme Inactivation Δ sfcA

PEP carboxikinase Overexpression Δ pck

Phosphotransacetylase Inactivation Δ pta-ackA

Pyruvate formate lyase Inactivation Δ pflB

Pyruvate oxidase Inactivation Δ poxB

Threonine decarboxylase Inactivation Δ tdcD

van Heerden and Nicol Microbial Cell Factories 2013, 12:80 Page 2 of 10http://www.microbialcellfactories.com/content/12/1/80

production, while providing an overview of the currentpetroleum process. The most reported SA-producing spe-cies include Actinobacillus succinogenes, Mannheimiasucciniciproducens, Anaerobiospirillum succiniciproducensand various strains of modified E. coli [4,10]. Thesebacteria, in addition to recombinant Corynebacteriumglutamicum, have been identified by McKinlay et al. [5]as the most promising SA producers. Limited develop-ment on yeast strains has also been reported, with lowpH fermentation the main objective [2,12,13].The open literature reports that the highest yields are

achieved by modified E. coli [14-18]. These yields are inexcess of 1 g SA per g glucose consumed (the theoret-ical maximum is 1.12 g g-1). All studies employed batchor fed-batch fermentation, with some utilising initialaerobic growth of cells [10,18]. From a bioreaction en-gineering perspective, these promising strains requirefurther exploration. This will allow the development ofnovel fermentor designs to optimise productivity andproduct titre. Given the projected scale of SA produc-tion, these will be key factors in the economic viabilityof the biobased process. In this regard the kinetic be-haviour of the cells is important, and growth and main-tenance rates, as well as inhibition characteristics, needto be quantified. It is also important to quantify the dis-tribution of metabolic products as a function of thestate of the cell (growing or non-growing) as this willaffect the overall yield [19-21].In the laboratory, batch or fed-batch fermentors are

preferred over continuous fermentors. The reason forthis is mostly a practical one, as batch or fed-batch sys-tems are not subject to the experimental challenges ofcontinuous operation.. Successful batch/fed-batch-basedprocesses from the laboratory are often directly scaledup to commercial units in order to obtain the desiredproduct. Up-scaling entails numerous engineering chal-lenges, and optimality cannot be guaranteed since theprocess is essentially a larger copy of the laboratory unit.For small-scale fermentors (producing high-value prod-ucts) process optimality is typically not the main eco-nomic driver. However, large-scale fermentation (forbulk chemical production such as SA) is different, as theefficiency of production is crucial to the economic suc-cess of the process. Continuous production has nu-merous advantages over conventional batch/fed-batchproduction, where higher productivities reduce capitalexpenditure and prolonged operation reduces operat-ing costs. Major disadvantages include possible strainmutation over time and inflexible operating conditions.Also, infections can cause upsets in the operation ofthe whole plant. Continuous conditions are superiorfor studying metabolic flux distribution, where steadystate allows proper mass balances to quantify fluxes ac-curately [22,23]. Accordingly, continuous conditions

are ideally suited for studying the performance charac-teristics of the microorganism.In this study one of the promising high-yield E. coli

strains, KJ134, from the group of Prof. L. Ingram from theUniversity of Florida was investigated. A yield of 1 g g-1,with a final SA titre of 71 g L-1 using a minimal mediumin a batch fermentor has been reported. All genetic modi-fications of the bacterium are given in Table 1 [14,24].E. coli KJ134 is evaluated under chemostat and batch con-ditions while continuous CO2 (g) sparging was used forsupplying inorganic carbon. Growth and maintenancecharacteristics are quantified in terms of rate, inhibitionand product distribution in order to provide a basis foroptimal fermentor design.

ResultsContinuous fermentationsAll steady state results are reported in Table 2. Geneticstability was confirmed by repeating dilution rates at dif-ferent times during the fermentation (the longest fer-mentation lasted 29 days). Steady state was confirmedby ensuring a steady dosing rate of potassium hydroxide(KOH) in excess of 10 hours. Figure 1 presents the SAconcentration and productivity for all steady states as afunction of the dilution rate employed. A similarity be-tween the 20 and 50 g L-1 glucose feed fermentations isobserved, although the 50 g L-1 fermentations appear tobe inhibited by the high substrate concentration at a di-lution rate of 0.09 h-1. This is evident from the start of adecline in productivity between dilution rates of 0.07and 0.09 h-1. The maximum growth rate for the 20 g L-1

glucose fermentation occurred at approximately 0.225 h-1

(extrapolation to SA concentration of 0 g L-1 at washout).

Table 2 Steady state results from continuous fermentations

Dilutionrate

Effective glucose infeeda

Glucose ineffluent

SA ineffluent

AcA ineffluent

Form A ineffluent

Malic acid ineffluent

DCW MassBalanceb

SA/AcA YSPc YSX P

(h-1) (g L-1) (g L-1) (g L-1) (g L-1) (g L-1) (g L-1) (g L-1) (mol mol-1) (g g-1) (g g-1) (g L-1 h-1)

0.056 19.0 2.2 12.9 2.9 0.3 0 NDd NDd 2.3 0.77 ND 0.72

0.146 19.4 12.4 5.0 1.4 0.6 0.38 0.88 1.00 1.9 0.71 0.125 0.73

0.146 19.4 12.7 4.7 1.4 0.5 0.00 0.87 1.00 1.7 0.70 0.130 0.68

0.093 19.0 6.6 9.2 2.5 0.6 0.34 1.97 1.05 1.9 0.74 0.159 0.85

0.093 19.0 6.4 9.3 2.5 0.6 0.00 1.62 1.02 1.9 0.74 0.128 0.87

0.187 19.8 15.8 2.5 0.8 0.3 0.51 0.68 1.00 1.5 0.62 0.171 0.46

0.037 19.0 0.0 14.0 3.9 0.5 0 NDd NDd 1.8 0.74 NDd 0.52

0.060 49.4 33.9 11.9 3.0 0.5 0.44 1.21 0.99 2.0 0.77 0.078 0.71

0.060 49.2 33.9 11.5 3.0 0.3 0.00 1.21 1.00 2.0 0.75 0.079 0.69

0.093 48.9 39.0 7.0 1.8 0.3 0.37 0.76 0.98 2.0 0.70 0.077 0.65

0.093 49.0 39.6 6.3 1.7 0.4 0.37 1.01 0.98 1.9 0.67 0.107 0.59

0.043 47.9 30.3 13.4 3.5 0.3 0.00 1.41 0.99 2.0 0.76 0.080 0.58

0.043 48.3 30.7 13.3 3.6 0.3 0.54 1.41 0.99 1.9 0.76 0.081 0.57

0.021 47.3 23.5 18.0 4.7 0.8 0.49 1.45 0.98 2.0 0.76 0.061 0.38

0.021 47.3 22.8 17.3 4.4 0.3 0.60 1.53 0.94 2.0 0.71 0.063 0.36a The dilution effect of KOH dosing incorporated into an effective feed concentration.b Fraction of accounted mass in effluent (malic acid not used in calculation).c SA yield based on glucose consumed.d Not determined.

vanHeerden

andNicolM

icrobialCellFactories2013,12:80

Page3of

10http://w

ww.m

icrobialcellfactories.com/content/12/1/80

0.0

0.2

0.4

0.6

0.8

1.0

1.2

0.00 0.05 0.10 0.15 0.20

Yie

ld (

g g-1

)

D (h-1)

YSP (50 g L-1 Glu)

YSP (20 g L-1 Glu)

YSX (50 g L-1 Glu)

YSX (20 g L-1 Glu)

Figure 2 Yield coefficients (YSP and YSX) from chemostatfermentations with glucose feed concentrations of 20 and50 g L-1. Glu – glucose.

0.0

1.0

2.0

3.0

4.0

0

4

8

12

16

20

0.00 0.05 0.10 0.15 0.20 0.25

Prod

uctiv

ity (

g L

-1h-1

)

Succ

inic

aci

d (g

L-1

)

D (h-1)

SA (20 g L-1 Glu)

SA (50 g L-1 Glu)

PSA (20 g L-1 Glu)

PSA (50 g L-1 Glu)

Figure 1 Steady state results from chemostat fermentationswith glucose feed concentrations of 20 and 50 g L-1. SA effluentconcentration and SA productivity given as a function of dilutionrate. Glu – glucose; SA – succinic acid.


The extrapolated value for the 50 g L-1 glucose fermen-tations appears to be slightly less.The SA concentration trend in Figure 1 suggests se-

vere product inhibition. At the lowest practically possibledilution rate of 0.02 h-1, the highest SA concentrationobtained was 18 g L-1 for the 50 g L-1 glucose (feed) con-centration. Glucose limitations prevented SA concentra-tions higher than 14 g L-1 for the 20 g L-1 glucose feedfermentation. The glucose residual at the extreme low di-lution of 0.02 h-1 suggest that growth terminates aroundthe observed steady state value of 18 g L-1 SA. Mainten-ance production by itself is not possible in a chemostatsince outlet biomass cannot be replaced by growth.Growth inhibition is likely caused by all the organic acids,but the available data do not allow distinction between theseparate effects of the different acids.Figure 2 gives the SA and biomass yields of the con-

tinuous fermentations. The SA yields are all between 0.7and 0.8 g g-1, except at the highest dilution rate wherelow glucose conversions complicate accurate yield deter-mination. Note from Table 2 that the SA to AcA ratio isclose to the value of two and that the formic acidamounts are small and considered negligible. The massbalances reported in Table 2 indicate that the measuredoutlet concentrations account reasonably well for the in-let mass flux, implying accuracy of the reported feed andproduct measurements.

Batch fermentationsThe results from the batch fermentation employing aninitial glucose concentration of 90 g L-1 are given in

Table 3, with the SA, AcA and dry cell weight (DCW)plotted in Figure 3. An additional batch fermentationwas performed with a starting glucose concentration of50 g L-1 (not reported). Very similar initial concentrationprofiles as that obtained with the 90 g L-1 initial glucosefermentation were observed. From Figure 3 the termin-ation of cell growth is observed at an SA concentrationof approximately 18 g L-1 and an AcA concentration of3.5 g L-1. This is similar to the continuous culture wheregrowth terminated around 18 g L-1 SA and 4.5 g L-1

AcA.Quantification of the SA yield on glucose is compli-

cated by the dilution effect of NaOH dosing and thecontinuous removal of samples from the system. A totalof 210 mL of 10 M KOH was dosed during the fermen-tation, and a total of 200 mL of broth was removed asseparate samples. The effective amount of glucose avail-able for the fermentation was estimated by performing atheoretical mass balance on each sample (see Table 3).Malic acid was emitted from the mass balance, due tothe difficulty of separating the malic acid peak duringthe HPLC (high-performance liquid chromatography)analysis from the glucose peak when the glucose con-centration was high. It is evident that the SA yield startsat a value similar to that of the continuous fermentations(below 0.8 g g-1) and thereafter increases to higher yieldvalues (up to 0.94 g g-1).The yield increase is also evident in the SA to AcA ra-

tio. Initially the ratio is similar to that of the continuousfermentations, but it already exceeds two while cellgrowth is still occurring. In order to analyse the

Table 3 Results of batch fermentation with an initial glucose concentration of 90 g L-1

Time Glucose SA AcA Form A DCW Glucose availablea YSPb SA/AcA ΔSA/ΔAcA

(h) (g L-1) (g L-1) (g L-1) (g L-1) (g L-1) (g L-1) (g g-1) (mol mol-1) (g g-1)

0.0 90.0 0.0 0.0 0.0 0.07 90 - - -

37.5 88.0 1.9 0.4 0.1 NDc 90 0.89 2.41 2.4

47.5 87.7 2.7 0.6 0.0 0.35 91 0.79 2.29 2.0

51.5 86.2 3.7 0.8 0.0 0.80 91 0.73 2.32 2.4

67.3 75.0 11.3 2.1 0.2 1.70 89 0.79 2.77 3.0

72.0 64.0 18.7 3.2 0.2 2.18 87 0.83 2.96 3.3

80.0 49.8 28.1 4.5 0.3 2.19 82 0.87 3.18 3.7

90.0 37.3 38.5 6.1 0.5 2.05 81 0.88 3.22 3.3

109.5 20.8 49.4 7.5 0.3 NDc NDc NDc 3.35 3.9

119.5 16.7 53.9 7.9 0.2 1.63 75 0.93 3.45 5.1

134.5 13.9 56.2 8.2 0.2 1.45 74 0.94 3.48 4.5

145.0 12.9 57.3 8.3 0.2 1.47 74 0.94 3.50 4.5

160.5 11.8 57.6 8.3 0.2 1.39 73 0.94 3.52

214.5 9.8 61.0 8.7 0.2 1.64 75 0.94 3.55 4.6a Glucose present and glucose converted to product (from mass balance).b Based on calculated glucose available.c Not determined.


instantaneous production characteristics of the cells, theincremental production ratio of SA to AcA should be con-sidered (see Table 3). At the point of growth terminationthe incremental ratio is three, while this value increases toas high as five during the maintenance production phase.The reason for the decrease in DCW in the mainten-

ance production phase is unclear, but was also observedin the (initial) 50 g L-1 glucose batch fermentation. The

0

1

2

3

4

0

10

20

30

40

50

60

70

0 50 100 150 200 250

DC

W (

g L

-1)

SA/A

cA c

once

ntra

tion

(g

L- 1

)

Time (h)

SA

AcA

DCW

Figure 3 Concentration profiles of the batch fermentation withan initial glucose concentration of 90 g L-1. SA – succinic acid;AcA – acetic acid; DCW – dry cell weight.

dilution effect due to sample removal and NaOHaddition cannot account for the drastic decrease in thebiomass content. It is also evident that the maintenanceproduction towards the end of the fermentation is ex-tremely slow, with significant amounts of residual glu-cose present (10 g L-1).

DiscussionSuccinic acid yield considerationsIn order to interpret the measured product distributionsand yields, proper analysis of the possible metabolicpathways is required. From an overall perspective theoptimum stoichiometry is given by the following bal-anced reaction where glucose (Glu) reacts with carbondioxide to form succinic acid (SA):

Gluþ 67CO2→

127SAþ 6

7H2O ð1Þ

The above equation ignores the formation of biomassand gives the maximum possible yield of SA based onone mol of glucose. When expressed in mass units theabove equation results in a SA yield of 1.12 g g-1.In order to understand metabolic pathways for

high-yield SA production, it is useful to consider thepathways of the naturally producing organisms such asActinobacillus succinogenes, Mannheimia succiniciproducensand Anaerobiospirillum succiniciproducens. All of these or-ganisms produce SA anaerobically via the reverse TCA(tricarboxylic acid) cycle starting with a PEP (phospho-enolpyruvate) carboxylation step. The net stoichiometryfrom glucose to SA requires NADH (see red arrows in


Figure 4) and the pathway is therefore referred to as thereductive pathway. In order to balance the redox an oxida-tive pathway is required (also starting at glucose as indi-cated by the blue arrows in Figure 4). All the wild SAproducers use AcA as a product to generate the requiredNADH. Pyruvate can be oxidised via the pyruvate de-hydrogenase (PDH) pathway or the formate lyase pathway(FL). In the PDH pathway NADH is released and accord-ingly the oxidative pathway in Figure 4(B) supplies morereduction power. When balancing the NADH of the twopathways the net oxidative and reductive flux can be de-termined. The combined overall pathway can then be usedto determine the maximum possible yield. For the FLroute in Figure 4(A) an equal amount of oxidative and re-ductive flux occur, implying a 1:1 split of carbon at thePEP node. The maximum mass-based yield for this sce-nario is 0.66 g g-1. For the PFL pathway (Figure 4(B)) theadditional NADH formed causes a ratio of 2:1 betweenthe reductive and oxidative paths, thus allowing for moreflux to SA where the maximum mass-based yield increasesto 0.87 g g-1. For this scenario each mol of AcA formedwill result in 2 mol of SA formed.It is possible to improve the metabolic pathways of the

natural organisms to further enhance SA yields by usingan oxidative pathway with SA as the final product. Thesimplest of these is the oxidative branch of the TCA cycle

Figure 4 Metabolic pathways of native SA producers. A) Pyruvate oxiddehydrogenase. Glu – glucose; PEP – phosphoenolpyruvate; OXA – oxaloacAcCoA – acetyl coenzyme A; AcA – acetate.

where reduction power or NADH is generated en route toSA (see Figure 5(A)). The TCA cycle can function in theabsence of oxygen, although no NADH can be convertedto ATP via oxidative phosphorylation. The major gener-ation of NADH in the oxidative side of the TCA cycle (seeFigure 5(A)) causes most of the flux to flow through thereductive part of the cycle (fraction of 5/7 of total flux).It is important to note that the combined stoichiometryis exactly the same as Equation 1. The representation inFigure 5(A) does not dictate termination of the oxidativeTCA section at SA. It is also possible to operate the fullTCA cycle as the oxidative pathway. In this case thereductive pathway flux will increase while the overallstoichiometry will remain the same as that of Equation 1.The glyoxylate shunt can be employed in a similar mannerto achieve the exact same result. The overall oxidativeroute will share sections of the reductive route. Thesmaller amount of NADH generated in the oxidativeglyoxylate path causes more oxidative flux (fraction of 3/7of total flux), as indicated in Figure 5(B).

Analysis of production capabilities and limitationsIt is evident that high-yield production of SA by E. coliKJ134 occurs under non-growth conditions where AcAproduction is small. The amount of AcA producedunder growth conditions, especially for continuous

ation via pyruvate formate lyase, B) Pyruvate oxidation via pyruvateetate; Mal – malate; Fum – fumarate; SA – succinate; Pyr – pyruvate;

Figure 5 Metabolic pathway with SA as the only excretion product. A) Reductive and oxidative TCA sections, B) Glyoxylate shunt.Glu – glucose; PEP – phosphoenolpyruvate; OXA – oxaloacetate; Mal – malate; Fum – fumarate; SA – succinate; Pyr – pyruvate; AcCoA – acetylcoenzyme A; Cit – citric acid; AcA – acetate.


cultures, resembles that of a native SA produceremploying pyruvate dehydrogenase for the oxidation ofpyruvate. This is clear for the SA:AcA ratios reportedin Table 2. All values are close to the theoretical valueof two (see Figure 4(B)) while formic acid productionis limited. The chemostat yield values all lie below thetheoretical maximum yield of 0.87 g g-1 due to biomassformation. The mechanism by which the AcA is pro-duced is unknown, but it is unlikely that it is due togenetic regression. This is because the chemostat fer-mentations exhibited good repeatability between initialand final steady states after extended operation. For thedevelopment of E. coli KJ134, numerous minor acetate-forming pathways were deleted [14], with minimal AcAformation reported in the test fermentations.Product distributions under non-growth or maintenance

conditions differ substantially from growth conditions asobserved in the batch results. It appears that the metabolicmodifications only function after cell growth has termi-nated, although SA:AcA ratios slightly higher than twowere obtained in the growth phase of the batch reactor.The highest SA:AcA ratio of five was obtained in themaintenance phase of the batch reactor, as seen from theinstantaneous production ratios in Table 3. It is thereforeevident that the ‘ideal’ metabolic pathway represented inFigure 5 is partially utilised under maintenance conditions,

where reduction power is generated via the oxidative TCA(or glyoxylate shunt) pathway.When comparing the batch runs of this study to that of

the original [1], the major difference lies in the inorganiccarbon source. Jantama et al. [14] employed a mixture ofK2CO3 and KOH for neutralisation and carbonate supply,while continuous CO2 (g) sparging was used in this study.Despite a major difference in the initial DCW (0.07 g L-1

for this study compared to 0.003 g L-1 previously repor-ted), the SA concentration after 96 hours was much higher(71.5 g L-1) than in this work (approximately 43 g L-1).The SA to AcA ratio was also significantly higher (16)than in this study (4).The growth inhibition characteristics of the chemostat

and batch fermentations agree, where a SA concentra-tion of 18 g L-1 (with the associated AcA) can be used asa marker. High-titre production of SA will accordinglyonly be possible with non-growing cells. The classicalfed-batch system is therefore suited for the process.High-titre continuous production will only be viablewith a cell recycle system where a separate growth fer-mentor supplies cells to a cell recycle fermentor. For theproposed scheme AcA production under growth condi-tions can be bypassed by operating the growth fermentoraerobically. The high SA titre from the non-growingcells will be associated with high SA yields.


Maintenance production rates were quantified for boththe chemostat and batch fermentations. The analysis ofthe continuous fermentations is given in Figure 6 wherethe cell-based production rate of SA (rSA) is plottedagainst the dilution rate. The y-axis intersection on thisgraph gives the maintenance production rate, which wasdetermined to be 0.3 g SA g-1 cells h-1. The maintenanceproduction rate for the batch fermentation can be ob-tained by dividing the slope of the SA profile (Figure 3)by the DCW (once there is no further increase in theDCW). It is evident that the slope of the SA profile de-creases towards the end of the fermentation. The max-imum slope (and maximum maintenance productionrate) was obtained right after growth termination, andwas determined to be 0.6 g SA g-1 cells h-1. This is twicethe rate of the chemostat estimate. The consequent de-crease in SA production could not be attributed to thedecrease in the DCW, implying that the rSA decreasestowards the end of the fermentation. Cell death mightbe a possible explanation for the observation.

ConclusionsTests on the modified E. coli KJ134 revealed that theflux distribution of the cell is not constant: higher SAfluxes were obtained under non-growth conditions, andSA yields in excess of 0.9 g g-1 are possible. Results fromthe continuous culture fermentations suggest that theoptimal SA-producing pathway is absent during growth,with AcA production supplying the reduction require-ments, similar to those of the wild SA producers. Theintended metabolic pathway, where the TCA oxidative

rSA = 3.06 D + 0.29

0.0

0.2

0.4

0.6

0.8

1.0

1.2

1.4

0.00 0.05 0.10 0.15 0.20

r SA

(g g

-1h-1

)

D (h-1)

SA (20 g L-1 Glu)

SA (50 g L-1 Glu)

Figure 6 rSA vs. D for the determination of the maintenanceproduction rate of SA. Glucose feed concentrations of 20 and50 g L-1 are indicated. Glu – glucose; SA – succinic acid.

branch (or the glyoxylate shunt) supplies the reductionrequirements, mainly functions under non-growth con-ditions, as observed from the batch fermentation.Growth termination occurred close to a SA concentra-

tion of 18 g L-1 – the amount of AcA from the batch andchemostat fermentations was similar. This indicated thathigh titre (as well as high yield) production is only possiblewith non-growing cells. The maximum maintenance ratewith SA production in the batch reactor was found to betwice that of the chemostat estimate.

MethodsMicroorganismThe modified E. coli KJ134 strain was supplied by the De-partment of Microbiology and Cell Science of the Univer-sity of Florida. The organism was genetically modified bythe research group of Prof. L. Ingram [14]. Vialscontaining treated beads in a cryopreservative solutionwere used to store culture samples at -75°C [25]. Allchemicals used in the fermentations were obtained fromMerck KgaA (Darmstadt, Germany) unless otherwise spe-cified. Temporary stock cultures were grown in Luria-Burtani broth. A defined medium named AM1, which wasdeveloped by Martinez et al. [26], was used in the fermen-tations and to grow the inoculum. The medium wassupplemented with 50 g L-1 D-glucose, 100 mM KHCO3

and 100 mM MOPS before stock culture from the Luria-Burtani broth was added. It was then incubated at 37°Cand 100 rpm over a period of 16–20 h in 30 mL sealedbottles containing 15 mL of the medium before the con-tinuous and batch reactors were inoculated.

Media composition20 or 50 g L-1 glucose was added to the AM1 medium(excluding betaine hydrochloride (HCl)) for the continu-ous fermentations. Batch fermentations were carried outwith initial glucose concentrations of 50 and 90 g L-1.Betaine HCl was used in the fermentation containing 90g L-1 glucose, as recommended by Martinez et al. [26]for media containing glucose concentrations in excess of50 g L-1. CO2 (g) (African Oxygen, Johannesburg, SouthAfrica) was used as the inorganic carbon source.

Continuous fermentationsThe bioreactor setup employed an external recycle foragitation and is described in detail by van Heerden et al.[27]. Each reservoir was fitted with a 0.2 μm PTFE (poly-tetrafluoroethylene) membrane filter (Midisart 2000 fil-ters from Sartorius, Göttingen, Germany). The workingvolume of the fermentations was 156 mL. The bioreactorconsisted of an aluminium top and bottom section and aglass tube with a length of 115 mm and an inner diam-eter of 37.5 mm. The bottom section contained oneentry for fermentation broth and also transferred heat


from the hotplate to the broth. The top section containeda sealed aluminium sheath that acted as a thermowell forthe thermocouple and two additional entry/exit points forthe broth. The complete setup was autoclaved for 40min at 121°C prior to starting the fermentation. The so-lution containing glucose and MgSO4 was autoclavedseparately from the solution containing the rest of thesalts to prevent precipitation and caramelisation of thesugar. The reactor was filled and operated at atemperature of 37°C and a pH of 7.00 ± 0.05 before itwas seeded with approximately 8–10 mL of the inocu-lum. 10 M non-sterile potassium hydroxide (KOH) wasused to control the pH. After startup the fermentor wasoperated at a dilution rate of 0.056 h-1 or lower. Thiswas done to approach batch conditions to allow for ini-tial accumulation of biomass. Continuous CO2 (g) flowinto the recycle line of the reactor was controlled at 5–10 mL min-1 (approximately 0.05 vvm) with a Brooksthermal mass flow controller. Antifoam A (Sigma-Al-drich, St. Louis, USA) was added to the fermentationbroth to control foam formation. Dilution rates between0.021 and 0.187 h-1 were employed and run continu-ously in excess of five turnovers to approach steadystate conditions. The stability and reproducibility of re-sults were investigated by repeating dilution rates at dif-ferent times in a fermentation and in fermentationsstarted up separately.

Batch fermentationsA 2 L Jupiter 2.0 fermentor (Solaris Biotechnology,Mantova, Italy) was used for the batch fermentations. Aworking volume of 1.5 L was used and operated at atemperature of 37°C and a pH of 7.00 ± 0.1. 10 M KOHwas used for pH control. 50 mL of inoculum wasinjected into the reactor after the temperature and pHhad stabilised. Antifoam A, diluted 1:9 with distilledwater, was pumped into the fermentor as required. CO2

(g) was continuously sparged into the fermentor at ap-proximately 0.3 L min-1 (approximately 0.2 vvm) beforeand during fermentations. Samples of 10–20 mL wereextracted aseptically from the reactor.

Analytical methodsGlucose and organic acid concentrations were determinedby using high-performance liquid chromatography. AnAgilent 1260 Infinity HPLC (Agilent Technologies, USA),equipped with an RI detector and a 300 × 7.8 mm AminexHPX-87H ion-exchange column (Bio-Rad Laboratories,USA), was used for this purpose. The mobile phase (0.3mL L-1 H2SO4) was fed at a flowrate of 0.6 mL min-1 withthe column temperature at 60°C. The DCW was deter-mined from 4.5–20 mL samples centrifuged at 750 g for10 min. The cell pellets were washed twice with distilledwater and dried at 90°C for at least 24 h.

AbbreviationsAcA: Acetic acid; D: Dilution rate (h-1); DCW: Dry cell weight; FL: Formatelyase; Glu: Glucose; HCl: Hydrochloride; KOH: Potassium hydroxide;PDH: Pyruvate dehydrogenase; PEP: Phosphoenolpyruvate; PSA: Productivity(g L-1 h-1); rSA: Specific maintenance production rate of succinic acid (g g-1

h-1); SA: Succinic acid; TCA: Tricarboxylic acid; YSP: Succinic acid yield (g g-1);YSX: Cell yield (g g-1).

Competing interestThe authors declare that they have no competing interests.

Authors’ contributionsCDVH and WN contributed to the design of the experiments. CDVHexecuted the fermentations and other experimental work. WN did theinterpretation, analysis and discussion of data. Both authors contributedtowards writing the manuscript.

Received: 7 August 2013 Accepted: 15 September 2013Published: 17 September 2013

References1. Bomgardner MM: Myriant to build succinic acid plant in Louisiana. Chem

Eng News 2011, 89:7.2. Yuzbashev TV, Yuzbasheva EY, Laptev IA, Sobolevskaya TI, Vybornaya TV,

Larina AS, Gvilava IT, Antonova SV, Sineoky SP: Is it possible to producesuccinic acid at a low pH? Bioeng Bugs 2011, 2:115–119.

3. Zeikus JG, Jain MK, Elankovan P: Biotechnology of succinic acidproduction and markets for derived industrial products. Appl MicrobiolBiotechnol 1999, 51:545–552.

4. Song H, Lee SY: Production of succinic acid by bacterial fermentation.Enzyme Microb Technol 2006, 39:352–361.

5. McKinlay JB, Vieille C, Zeikus JG: Prospects for a bio-based succinateindustry. Appl Microbiol Biotechnol 2007, 76:727–740.

6. Bechthold I, Bretz K, Kabasci S, Kopitzky R, Springer A: Succinic acid: A newplatform chemical for biobased polymers from renewable resources.Chem Eng Technol 2008, 31:647–654.

7. Cukalovic A, Stevens CV: Feasibility of production methods for succinicacid derivatives: a marriage of renewable resources and chemicaltechnology. Biofuels Bioprod Biorefin 2008, 2:505–529.

8. Delhomme C, Weuster-Botz D, Kühn FE: Succinic acid from renewableresources as a C4 building-block chemical—a review of the catalyticpossibilities in aqueous media. Green Chem 2009, 11:13–26.

9. Chimirri F, Bosco F, Ceccarelli R, Venturello A, Geobaldo F: Succinic acidand its derivatives : fermentative production using sustainable industrialagro-food by-products and its applications in the food industry. Ital JFood Sci 2010, 22:119–125.

10. Beauprez JJ, De Mey M, Soetaert WK: Microbial succinic acid production:Natural versus metabolic engineered producers. Process Biochem 2010,45:1103–1114.

11. Cheng KK, Wang GY, Zeng J, Zhang JA: Improved succinate production bymetabolic engineering. BioMed Res Int 2013:1–12.

12. Raab AM, Lang C: Oxidative versus reductive succinic acid production inthe yeast Saccharomyces cerevisiae. Bioeng Bugs 2011, 2:120–123.

13. Smidt M: A sustainable supply of succinic acid. Euro Biotech News 2011,10(C):11–12.

14. Jantama K, Zhang X, Moore JC, Shanmugam KT, Svoronos SA, Ingram LO:Eliminating side products and increasing succinate yields in engineeredstrains of Escherichia coli C. Biotechnol Bioeng 2008, 101:881–893.

15. Hong SH, Lee SY: Importance of redox balance on the production ofsuccinic acid by metabolically engineered Escherichia coli. Appl MicrobiolBiotechnol 2002, 58:286–290.

16. Vemuri GN, Eiteman MA, Altman E: Succinate production in dual-phaseEscherichia coli fermentations depends on the time of transition fromaerobic to anaerobic conditions. J Ind Microbiol Biotechnol 2002,28:325–332.

17. Isar J, Agarwal L, Saran S, Saxena RK: A statistical method for enhancingthe production of succinic acid from Escherichia coli under anaerobicconditions. Bioresource Technol 2006, 97:1443–1448.

18. Balzer GJ, Thakker C, Bennett GN, San KY: Metabolic engineering ofEscherichia coli to minimize byproduct formate and improving succinateproductivity through increasing NADH availability by heterologous


expression of NAD(+)-dependent formate dehydrogenase. Metab Eng2013, 20:1–8.

19. Leib TM, Pereira CJ, Villadsen J: Bioreactors: a chemical engineeringperspective. Chem Eng Sci 2001, 56:5485–5497.

20. Lübbert A, Jørgensen SB: Bioreactor performance: a more scientificapproach for practice. J Biotechnol 2001, 85:187–212.

21. Villadsen J: Innovative technology to meet the demands of the whitebiotechnology revolution of chemical production. Chem Eng Sci 2007,62:6957–6968.

22. Shuler M, Kargi F: Bioprocess Engineering. Second Edi. New Jersey: PrenticeHall; 2002.

23. Villadsen J, Nielsen J, Lidén G, Principles BE: Bioreaction EngineeringPrinciples. Third Edit. New York: Springer; 2011.

24. Jantama K, Haupt MJ, Svoronos SA, Zhang X, Moore JC, Shanmugam KT,Ingram LO: Combining metabolic engineering and metabolic evolutionto develop nonrecombinant strains of Escherichia coli C that producesuccinate and malate. Biotechnol Bioeng 2008, 99:1140–1153.

25. Pro-Lab Diagnostics: Microbank™. Neston, UK: Worldwide PerformancePortfolio; 2011:1–78.

26. Martinez A, Grabar TB, Shanmugam KT, Yomano LP, York SW, Ingram LO:Low salt medium for lactate and ethanol production by recombinantEscherichia coli B. Biotechnol Lett 2007, 29:397–404.

27. Van Heerden CD, Nicol W: Continuous succinic acid fermentation byActinobacillus succinogenes. Biochem Eng J 2013, 73:5–11.

doi:10.1186/1475-2859-12-80Cite this article as: van Heerden and Nicol: Continuous and batchcultures of Escherichia coli KJ134 for succinic acid fermentation:metabolic flux distributions and production characteristics. Microbial CellFactories 2013 12:80.

Submit your next manuscript to BioMed Centraland take full advantage of:

• Convenient online submission

• Thorough peer review

• No space constraints or color figure charges

• Immediate publication on acceptance

• Inclusion in PubMed, CAS, Scopus and Google Scholar

• Research which is freely available for redistribution

Submit your manuscript at www.biomedcentral.com/submit

Documents

Continuous and batch cultures of Escherichia coli KJ134